Process and system for the production of ethylene carbonate and/or ethylene glycol

ABSTRACT

The invention relates to a process for producing ethylene carbonate and/or ethylene glycol, which comprises the following steps: a) supplying an overhead absorber stream withdrawn from an absorber to a vapor-liquid separator to yield an aqueous bottoms stream and a recycle gas stream; b) supplying an aqueous process stream comprising one or more impurities to a distillation apparatus to yield an overhead impurities stream and a purified aqueous process stream, wherein the aqueous process stream supplied to the distillation apparatus comprises at least a portion of the aqueous bottoms stream withdrawn from the vapor-liquid separator, wherein the overhead impurities stream is supplied to a condenser and is cooled to a temperature in the range of from 5 to 95° C., wherein the cooled overhead impurities stream is split into a reflux stream which is recycled to the distillation apparatus and an overhead impurities stream; and further steps c) and d).

FIELD OF THE INVENTION

The present invention relates to a process and a system for theproduction of ethylene carbonate and/or ethylene glycol.

BACKGROUND OF THE INVENTION

Ethylene glycol is a valuable industrial compound that is widelyemployed as starting material for the manufacture of polyester fibresand polyethylene terephthalate (PET) resins. It also finds applicationin automotive antifreeze and hydraulic brake fluids, aircraft de-icersas well as in pharmaceutical products.

Ethylene glycol is typically prepared from ethylene oxide, which is inturn prepared by the silver-catalyzed oxidation of ethylene. Morespecifically, ethylene and oxygen are passed over a silver-basedepoxidation catalyst, typically at pressures of 10-30 bar andtemperatures of 200-300° C., producing a product stream comprisingethylene oxide, carbon dioxide, ethylene, oxygen and water.

In one well-known process, ethylene oxide is then reacted with a largeexcess of water in a non-catalytic process, producing a glycol productstream comprising close to 90 wt. % monoethylene glycol (MEG), theremainder being predominantly diethylene glycol (DEG), some triethyleneglycol (TEG) and a small amount of higher homologues.

Further, in another well-known process, ethylene oxide is reacted withcarbon dioxide in the presence of a catalyst to produce ethylenecarbonate, which is subsequently hydrolyzed to provide ethylene glycol.Reaction via ethylene carbonate significantly improves the selectivityof ethylene oxide conversion to monoethylene glycol.

Still further, in another known process (see e.g. EP2178815),monoethylene glycol is prepared in a reactive absorption process,wherein the gases from the ethylene oxide reactor are supplied to areactive absorber and the ethylene oxide is contacted with an aqueouslean absorbent in the presence of one or more carboxylation andhydrolysis catalysts, and wherein the majority of the ethylene oxide isconverted to ethylene carbonate or ethylene glycol in the absorber.

In the above-mentioned reactive absorption process, considerable volumesof process water are produced. It is often desirable from an economicperspective to recycle as much of the process water as possible, forexample, by recycling the process water to the ethylene oxide reactiveabsorber (e.g. for use as lean absorbent). Advantageously, suchrecycling not only reduces operating costs because it reduces therequisite amount of fresh water to be supplied to the process, but itmay also reduce costs associated with disposal of the process water aswaste.

However, process water often contains various impurities, which areoften a result of the formation of by-products during the production ofethylene oxide, ethylene carbonate and/or ethylene glycol. For example,an overhead stream withdrawn from an ethylene oxide reactive absorbertypically comprises, in addition to water, hydrocarbon impurities suchas formaldehyde, acetaldehyde, etc. Additionally, organic chlorideimpurities may also be found in process water.

If process water is recycled to the ethylene oxide reactive absorberwithout first removing at least a portion of these impurities, then overtime, the impurities accumulate and are deleterious to the overallquality of the resulting glycol product and/or cause catalystdegradation.

WO2017178418 discloses an ethylene oxide reactive absorption process andsystem for the production of ethylene carbonate and/or ethylene glycol,wherein an aqueous process stream comprising one or more impurities issupplied to a distillation apparatus and wherein said distillationapparatus distills a majority of the impurities present in the aqueousprocess stream overhead and out of the process. Further, saidWO2017178418 discloses that said aqueous process stream may first besupplied to a heating mechanism (such as a preheater) and then to aflash vessel, before supplying it to the distillation apparatus.

As mentioned above, organic chloride impurities may also be found in theprocess water. This is caused by the use of organic chloride moderatorsin the epoxidation reaction, which moderators control the performance ofthe epoxidation catalyst. Commonly used moderators includemonochloroethane or dichloroethane. These organic chloride moderatorsare partially converted to other organic chloride compounds whichcomprise 2-chloroethanol and chloromethyl dioxolane. The presence ofthese organic chloride impurities can lead to problems. For theseorganic chloride compounds can be converted to ethylene oxide andinorganic chloride compounds (e.g. potassium chloride). Thus, theseorganic chloride compounds can lead to a build-up of inorganic chloride.The inorganic chloride in turn can start to precipitate and may causechloride stress corrosion.

Therefore, it is an object of the present invention to provide anethylene oxide reactive absorption process and system for the productionof ethylene carbonate and/or ethylene glycol, wherein both (i) organicchloride impurities (such as 2-chloroethanol and chloromethyl dioxolane)are removed from the process as much as possible and (ii) anaccumulation of inorganic chloride in that process through conversion ofsaid organic chloride impurities is prevented or minimized.

SUMMARY OF THE INVENTION

Surprisingly it was found that the object of maximizing organic chlorideremoval and minimizing inorganic chloride formation, as discussed above,may be achieved by a process as disclosed in above-mentionedWO2017178418 in which latter process an aqueous process streamcomprising one or more impurities is supplied to a distillationapparatus to yield an overhead impurities stream and a purified aqueousprocess stream, characterized in that in the process of the presentinvention the overhead impurities stream is supplied to a condenser andis cooled to a temperature in the range of from 5 to 95° C., wherein thecooled overhead impurities stream is split into a reflux stream which isrecycled to the distillation apparatus and an overhead impuritiesstream.

Accordingly, the present invention relates to a process for theproduction of ethylene carbonate and/or ethylene glycol comprising:

a) supplying an overhead absorber stream withdrawn from an absorber to avapor-liquid separator to yield an aqueous bottoms stream and a recyclegas stream;

b) supplying an aqueous process stream comprising one or more impuritiesto a distillation apparatus to yield an overhead impurities stream and apurified aqueous process stream, wherein the aqueous process streamsupplied to the distillation apparatus comprises at least a portion ofthe aqueous bottoms stream withdrawn from the vapor-liquid separator,wherein the overhead impurities stream is supplied to a condenser and iscooled to a temperature in the range of from 5 to 95° C., wherein thecooled overhead impurities stream is split into a reflux stream which isrecycled to the distillation apparatus and an overhead impuritiesstream;

c) supplying at least a portion of the purified aqueous process streamand an ethylene oxide product stream to the absorber; and

d) contacting the ethylene oxide product stream with the purifiedaqueous process stream in the absorber in the presence of one or morecarboxylation and hydrolysis catalysts to yield a fat absorbent streamcomprising ethylene carbonate and/or ethylene glycol.

Further, the present invention relates to a system for the production ofethylene carbonate and/or ethylene glycol, wherein the system comprises:

an absorber comprising one or more carboxylation and hydrolysiscatalysts, at least two inlets and at least two outlets, wherein a firstinlet of the absorber is fluidly connected to an outlet of an ethyleneoxide reactor;

a vapor-liquid separator comprising an inlet and an outlet, wherein theinlet of the vapor-liquid separator is fluidly connected to a firstoutlet of the absorber;

a distillation apparatus comprising an inlet and two outlets, whereinone outlet of the distillation apparatus is fluidly connected to asecond inlet of the absorber and the other outlet of the distillationapparatus is fluidly connected to a sub-cooling condenser; and

optionally a preheater comprising an inlet and an outlet, wherein theinlet of the preheater is fluidly connected to the outlet of thevapor-liquid separator and wherein the outlet of the preheater isfluidly connected to the inlet of the distillation apparatus.

BRIEF DESCRIPTION OF THE DRAWINGS

FIGS. 1 and 2 are schematic illustrations showing exemplary embodimentsof the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The process of the present invention comprises various steps, asdescribed hereinbelow. Said process may comprise one or moreintermediate steps between these steps. Further, said process maycomprise one or more additional steps preceding the first mentioned stepand/or following the last mentioned step.

While the process and system of the present invention and mixtures orstreams used in said process are described in terms of “comprising”,“containing” or “including” one or more various described steps orcomponents, they can also “consist essentially of” or “consist of” saidone or more various described steps or components.

In the context of the present invention, in a case where a mixture,stream or catalyst comprises two or more components, these componentsare to be selected in an overall amount not to exceed 100%.

Within the present specification, “substantially no” means that nodetectible amount of the component in question is present in thecatalyst or composition.

Further, where upper and lower limits are quoted for a property then arange of values defined by a combination of any of the upper limits withany of the lower limits is also implied.

Described herein are processes and systems for the production ofethylene carbonate and/or ethylene glycol. By using the processes andsystems disclosed herein, it is possible to reduce the amount ofimpurities present in an aqueous process stream, thereby allowing all orsubstantially all of the process water generated in the production ofethylene glycol to be reused. More specifically, when an aqueous processstream comprising one or more impurities is supplied to a distillationapparatus in accordance with the present disclosure, the distillationapparatus distills a majority of the impurities present in the aqueousprocess stream overhead and out of the process, while water and valuableglycol product is returned to the process as a purified aqueous processstream.

Accordingly, the systems and processes disclosed herein provide theadvantage that the volume of waste water generated in the production ofethylene carbonate and/or ethylene glycol may be reduced, and furtherthat impurities are not permitted to accumulate in the process and havea deleterious effect on the quality of the resulting glycol product.Additionally, by carrying out these processes in an ethylene glycolmanufacturing plant, it is possible to significantly reduce the amountof fresh water required, reduce the amount of waste water produced andreduce catalyst degradation, all of which reduce operating costs.

Specifically, in step b) of the present process, the overhead impuritiesstream from the distillation apparatus is supplied to a condenser and iscooled to a temperature in the range of from 5 to 95° C., preferably 10to 85° C., more preferably 20 to 75° C., more preferably 25 to 65° C.,most preferably 30 to 60° C., wherein the cooled overhead impuritiesstream is split into a reflux stream which is recycled to thedistillation apparatus and an overhead impurities stream which is notrecycled to the distillation apparatus.

Preferably, in the present invention, the above-mentioned condenser is aso-called “sub-cooling” condenser, which in the present specificationrefers to a condenser which is capable of cooling to a temperature belowthe condensation temperature, in specific cooling to a temperature inthe above-mentioned range of from 5 to 95° C.

Preferably, in step b) of the present process, the aqueous processstream comprising one or more impurities is supplied first to apreheater before it is supplied to a distillation apparatus. Further, itis preferred that said preheater is positioned at a minimum distance tothe inlet of the distillation apparatus. Within the presentspecification, the phrase “at a minimum distance” means that thedistance between the outlet of the preheater and the inlet of thedistillation apparatus is short. Preferably, said distance is as shortas possible as can be allowed by interconnecting pipe design. Saiddistance may for example be of from 0.5 to 10 metres, preferably 1 to 10metres, more preferably 1 to 5 metres, most preferably 1 to 3 metres.This implies that if there is a line (pipe) connecting the outlet of thepreheater to the inlet of the distillation apparatus, the length of thatline corresponds to said short distance. Still further, it is preferredthat said preheater has a minimum volume.

The process of producing ethylene glycol and/or ethylene carbonate byepoxidation of ethylene and reactive absorption of ethylene oxide hasbeen described in detail in, among others, WO2009021830, WO2009140318,WO2009140319, the disclosures of which are hereby incorporated byreference.

Typically, the epoxidation process comprises reacting, in an ethyleneoxide reactor, ethylene with oxygen in the presence of an epoxidationcatalyst to form ethylene oxide. In such a reaction, the oxygen may besupplied as oxygen or as air, but is preferably supplied as oxygen.Ballast gas, for example methane or nitrogen, is typically supplied toallow operation at high oxygen levels without causing a flammablemixture. Moderator, e.g. monochloroethane (ethyl chloride), vinylchloride or dichloroethane, may be supplied for ethylene oxide catalystperformance control.

The ethylene oxide reactor is typically a multitubular, fixed bedreactor. The epoxidation catalyst preferably comprises silver andoptionally promoter metals deposited on a support material, for example,alumina. The epoxidation reaction is preferably carried out at pressuresof greater than 1 MPa and less than 3 MPa and temperatures of greaterthan 200° C. and less than 300° C. The ethylene oxide product streamwithdrawn from the ethylene oxide reactor is preferably cooled in one ormore coolers, preferably with generation of steam at one or moretemperature levels.

The ethylene oxide product stream from the ethylene oxide reactor, whichtypically comprises ethylene oxide, unreacted reactants (i.e., ethyleneand oxygen), carbon dioxide, and water, is then passed to an absorber inwhich it is intimately contacted with lean absorbent. Typically, thelean absorbent comprises at least 20 wt % water, and preferablycomprises from 20 wt % to 80 wt % water. The lean absorbent may alsocomprise ethylene glycol.

In the absorber, the ethylene oxide product stream is intimatelycontacted with the lean absorbent in the presence of one or morecarboxylation and hydrolysis catalysts. If this occurs in the presenceof only one catalyst, then the catalyst must promote carboxylation andhydrolysis. If this occurs in the presence of two or more catalysts,then each catalyst can promote carboxylation or hydrolysis or canpromote both reactions (provided that at least one catalyst promotescarboxylation and at least one catalyst promotes hydrolysis).Preferably, the ethylene oxide product stream is contacted with leanabsorbent in the presence of at least two catalysts including a firstcatalyst that promotes carboxylation and a second catalyst that promoteshydrolysis. Suitably, the absorber may be the sort of reactive absorberdescribed in WO2009021830 or WO2016046100, the disclosures of which arehereby incorporated by reference.

Preferably, the one or more carboxylation and hydrolysis catalystsis/are homogenous, and the lean absorbent contains the one or morecatalysts. Homogeneous catalysts that are known to promote carboxylationinclude alkali metal halides such as potassium iodide and potassiumbromide, and halogenated organic phosphonium or ammonium salts such astributylmethylphosphonium iodide, tetrabutylphosphonium iodide,triphenylmethylphosphonium iodide, triphenyl-propylphosphonium bromide,triphenylbenzylphosphonium chloride, tetraethylammonium bromide,tetramethylammonium bromide, benzyltriethylammonium bromide,tetrabutylammonium bromide and tributylmethylammonium iodide. Preferredhomogeneous catalysts that are known to promote carboxylation includealkali metal iodides such as potassium iodide and halogenated organicphosphonium or ammonium salts such as tributylmethylphosphonium iodide,tetrabutylphosphonium iodide, triphenylmethylphosphonium iodide andtributylmethylammonium iodide.

Homogeneous catalysts that are known to promote hydrolysis include basicalkali metal salts such as potassium carbonate, potassium hydroxide andpotassium bicarbonate, or alkali metal metalates such as potassiummolybdate. Preferred homogeneous catalyst systems include a combinationof potassium iodide and potassium carbonate, and a combination ofpotassium iodide and potassium molybdate.

In another embodiment, the one or more carboxylation and hydrolysiscatalysts is/are heterogeneous and the heterogeneous catalyst(s) is/arecontained in vertically stacked trays. Heterogeneous catalysts thatpromote carboxylation include quaternary ammonium and quaternaryphosphonium halides immobilized on silica, quaternary ammonium andquaternary phosphonium halides bound to insoluble polystyrene beads, andmetal salts such as zinc salts immobilised on solid supports containingquaternary ammonium or quaternary phosphonium groups, such as ionexchange resins containing quaternary ammonium or quaternary phosphoniumgroups. Heterogeneous catalysts that promote hydrolysis includemetalates immobilised on solid supports, for example molybdates,vanadates or tungstates immobilised on ion exchange resins containingquaternary ammonium or quaternary phosphonium groups, or basic anionssuch as bicarbonate ions immobilised on solid supports, for examplebicarbonate immobilised on ion exchange resins containing quaternaryammonium or quaternary phosphonium groups.

The temperature in the absorber is preferably from 50° C. to 160° C.,preferably from 80° C. to 150° C., more preferably from 80° C. to 120°C. This is higher than the temperature in an absorber in a conventionalprocess and is required to promote the carboxylation and hydrolysisreactions. Temperature higher than 160° C. is not preferred as this mayreduce the selectivity of ethylene oxide conversion to ethylene glycol.Both the ethylene oxide product stream and the lean absorbent arepreferably supplied to the absorber at temperatures in the range from50° C. to 160° C.

The pressure in the absorber is from 1 to 4 MPa, preferably from 2 to 3MPa. The preferred pressure is a compromise between lower pressures thatrequire less expensive equipment (e.g. equipment having thinner walls)and higher pressures that increase absorption and reduce the volumetricflow of the gas, thereby reducing the size of equipment and piping.

At least 50% of the ethylene oxide entering the absorber may beconverted in the absorber. Preferably, at least 60%, more preferably atleast 70%, even more preferably at least 80%, most preferably at least90% of the ethylene oxide entering the absorber is converted inabsorber. The ethylene oxide may undergo carboxylation, providingethylene carbonate. The ethylene oxide may undergo hydrolysis, providingethylene glycol. Additionally, the ethylene carbonate that is producedfrom the ethylene oxide may undergo hydrolysis, providing ethyleneglycol.

The ethylene oxide product stream supplied to the absorber comprisescarbon dioxide. However, it is possible that the ethylene oxide productstream may contain insufficient carbon dioxide to achieve desired levelsof carboxylation. Optionally, an additional source of carbon dioxide issupplied to the absorber, e.g. recycle carbon dioxide from a finishingreactor, carbon dioxide from a carbon dioxide recovery unit or, atstart-up, carbon dioxide from an external source.

A ‘fat absorbent’ stream is withdrawn from the absorber, preferably bywithdrawing liquid from the bottom of the absorber. The fat absorbentstream will comprise ethylene carbonate and/or ethylene glycol and anyremaining ethylene oxide, if present, depending on the conditions,set-up and catalyst in the absorber. In addition, when the one or morecarboxylation and hydrolysis catalysts is/are homogenous, the fatabsorbent stream will further comprise the one or more carboxylation andhydrolysis catalysts.

Optionally, a portion or all of the fat absorbent stream is supplied toone or more finishing reactors (e.g., to provide further conversion ofany ethylene oxide and/or ethylene carbonate that was not converted toethylene glycol in the absorber). Suitable finishing reactors mayinclude a carboxylation reactor, a hydrolysis reactor, a carboxylationand hydrolysis reactor, and a combination thereof. Supply to one or morefinishing reactors is preferred if a significant quantity (e.g. at least1%) of ethylene oxide or ethylene carbonate is not converted to ethyleneglycol in the absorber. To maximize conversion of ethylene oxide in theabsorber, spraying nozzles can be employed in the sump (bottom section)of the absorber, to disperse carbon dioxide and promote carboxylation.Optionally, steam may be injected into a finishing reactor suitable forhydrolysis.

Carbon dioxide may be produced in the one or more finishing reactorsand, if desired, may be separated from the one or more finishing reactorproduct stream(s) as it leaves the one or more finishing reactors and isoptionally recycled to the absorber.

The temperature in the one or more finishing reactors is typically from100° C. to 200° C., preferably from 100° C. to 180° C. The pressure inthe one or more finishing reactors is typically from 0.1 to 3 MPa.

The fat absorbent stream or a finishing reactor product stream isoptionally supplied to a flash vessel or to a light ends stripper. Lightends are removed in the flash vessel or in the light ends stripper.(Light ends are gases such as ethylene, and also ballast gases such asmethane.) Optionally, if desired, flash vaporization may be achieved ina finishing reactor (e.g., hydrolysis reactor) so that a separate flashvessel may not be required and the equipment used in the process isthereby reduced.

Optionally, a flash vessel may be located directly after the absorber sothe fat absorbent stream passes directly from an outlet of the absorberto the flash vessel. When there is at least one finishing reactor, aflash vessel may be located after all of the one or more finishingreactors so that the finishing reactor product stream passes from saidfinishing reactors to the flash vessel. When there is more than onefinishing reactor, a flash vessel may be located between the finishingreactors such that the fat absorbent stream passes from the absorber toat least one finishing reactor, then the finishing reactor productstream passes to the flash vessel and then the stream from the flashvessel passes to at least another finishing reactor. The flash can be atpressure from 0.01 to 2 MPa, preferably from 0.1 to 1 MPa, mostpreferably from 0.1 to 0.5 MPa.

The fat absorbent stream from the absorber, or the finishing reactorproduct stream from one or more finishing reactors or other productstream comprising ethylene glycol is supplied to a dehydrator as adehydrator feed stream. The dehydrator feed stream preferably comprisesvery little ethylene oxide or ethylene carbonate, i.e. most of theethylene oxide or ethylene carbonate has been converted to ethyleneglycol prior to supply to the dehydrator, either in the absorber or in afinishing reactor. Preferably the molar ratio of ethylene glycol toethylene oxide and ethylene carbonate (combined) in the dehydrator feedstream is greater than 90:10, more preferably greater than 95:5, evenmore preferably greater than 99:1, and most preferably at 999:1.Suitably, the dehydrator feed stream may comprise 10 ppm or less ofethylene carbonate.

The dehydrator is preferably one or more columns, including at least onevacuum column, preferably operating at a pressure of less than 0.05 MPa,more preferably less than 0.025 MPa and most preferably about 0.0125MPa.

An overhead dehydrator stream, which generally comprises water and oneor more impurities, is withdrawn from the dehydrator, typically at ornear the top of the dehydrator. All or a portion of the overheaddehydrator stream is then supplied to the absorber, the distillationapparatus, or a combination thereof. For example, all or a portion ofthe overhead dehydrator stream may be combined with a purified aqueousprocess stream and supplied to the absorber, combined with an aqueousbottoms stream from a vapor liquid separator and supplied to thedistillation apparatus as an aqueous process stream, or a combinationthereof. Further, if desired, a portion of the overhead dehydratorstream may optionally be combined with an overhead impurities streamfrom the distillation apparatus and disposed of as waste.

A dehydrator bottom stream, comprising predominantly MEG, is withdrawnfrom the dehydrator, typically at or near the bottom of the dehydrator,and is optionally supplied to a separator (e.g., an evaporator or asplitter) and/or to a glycol purification apparatus (e.g., a glycolpurification column) to remove impurities. When a separator is used, aglycol product stream is withdrawn from the separator, typically at ornear the top, and is optionally further supplied to a glycolpurification apparatus. Further, a hot process stream (e.g., a catalystrecycle stream or a glycol absorbent stream) is withdrawn from theseparator, typically at or near the bottom of the separator, and isoptionally recycled to the absorber. In those embodiments where the oneor more carboxylation and hydrolysis catalysts used is/are homogeneouscatalysts, the one or more homogenous catalysts may be separated fromthe dehydrator bottom stream in the separator as a catalyst recyclestream and recycled to the absorber for reuse therein. Similarly, inthose embodiments where the one or more carboxylation and hydrolysiscatalysts used is/are heterogeneous catalysts, a glycol absorbent streammay be withdrawn from the separator and recycled to the absorber forreuse therein.

Suitably, a hot process stream, such as a catalyst recycle stream or aglycol absorbent stream, may be withdrawn from the separator, cooled andcombined with a purified aqueous process stream before being recycled tothe absorber. If desired, all or a portion of the heat removed from thehot process stream may be recovered and utilized via process heatintegration to provide the requisite thermal energy needed in otherparts of the process, as discussed further below.

Gases that are not absorbed in the absorber are removed at or near thetop of the absorber and condensed to yield an overhead absorber stream,which is supplied to a vapor-liquid separator, such as a knock-outvessel, flash vessel, etc. A recycle gas stream, which typicallycomprises unreacted reactants (e.g., ethylene and oxygen), ballast gas(e.g., methane), carbon dioxide, etc., is withdrawn from thevapor-liquid separator, typically at or near the top. Optionally, atleast a portion of the recycle gas stream withdrawn from thevapor-liquid separator is supplied to a carbon dioxide absorptioncolumn, wherein carbon dioxide is at least partially absorbed by arecirculating absorbent stream, and/or to one or more guard beds,wherein halogen-containing impurities may be at least partially absorbedby a purification absorbent, before being recycled to the ethylene oxidereactor. Suitably, the one or more guard beds may be the sort of guardbeds described in WO2017102694, WO2017102698, WO2017102701 andWO2017102706, the disclosures of which are hereby incorporated byreference.

An aqueous bottoms stream, which generally comprises water, one or moreimpurities and optionally glycols, is withdrawn from the vapor-liquidseparator, typically at or near the bottom, and at least a portion ofthe aqueous bottoms stream is then supplied to a distillation apparatusas an aqueous process stream. Optionally, if desired, a portion of theaqueous bottoms stream may bypass the distillation apparatus and becombined with a purified aqueous process stream withdrawn from thedistillation apparatus and supplied to the absorber.

In accordance with the present disclosure, an aqueous process streamcomprising one or more impurities is supplied to a distillationapparatus. As previously mentioned, the aqueous process stream suppliedto the distillation apparatus comprises at least a portion of theaqueous bottoms stream withdrawn from the vapor-liquid separator.Additionally, the aqueous process stream supplied to the distillationapparatus may comprise at least a portion of the overhead dehydratorstream withdrawn from the dehydrator. By supplying the aqueous processstream to the distillation apparatus, the amount of one or moreimpurities present therein is reduced via distillation.

Typically, the aqueous process stream supplied to the distillationapparatus comprises a major amount of water (i.e., an amount of watergreater than or equal to 88 wt. %, for example from about 89.5 to 99 wt.%, relative to the total weight of the aqueous process stream) and aminor amount of one or more impurities (i.e., a total amount ofimpurities of less than 0.6 wt. %, for example from about 0.1 to 0.5 wt.%, or from 0.2 to 0.4 wt. % relative to the total weight of the aqueousprocess stream). Optionally, the aqueous process stream furthercomprises glycols (e.g., monoethylene glycol (“MEG”)) in an amount up to12 wt. %, for example from about 0.5 to 10 wt. %, relative to the totalweight of the aqueous process stream.

Examples of one or more impurities that may be found in an aqueousprocess stream include, but are not necessarily limited to, hydrocarbonand chlorinated hydrocarbon impurities such as aldehydes, alcohols,acetals, cyclic acetals, ethers, cyclic ethers, and esters, for exampleformaldehyde, acetaldehyde, glycolaldehyde, propionaldehyde,2,3-epoxy-1,4-dioxane, 1,4-dioxane, 1,3-dioxolane,2-methyl-1,3-dioxolane, 2-methoxy ethanol, ethanol, 2-ethoxy ethanol,2-hydroxymethyl-1,3-dioxolane, 2,2′-bis-1,3-dioxolane,2-chloromethyl-1,3-dioxolane, hydroxyacetone, 2-chloroethanol,glycolates, formates, lactates, acetates, propionates and a combinationthereof.

Before being supplied to the distillation apparatus, the aqueous processstream may be supplied to a preheater. Further, optionally, before beingsupplied to said preheater, the aqueous process stream may be suppliedto a flash vessel to recover light ends (e.g., ethylene and methane),which are preferably recycled to the ethylene oxide reactor aftercompression. For example, in one embodiment, the aqueous process streamis first supplied to a flash vessel and then to the preheater beforebeing supplied to the distillation apparatus. In practice, the pressureof the flash vessel should be higher than the pressure of thedistillation apparatus. Therefore, the flash is typically at a pressureof from 100 kPa to 270 kPa, or from 130 kPa to 220 kPa, or from 170 kPato 210 kPa. Preferably, the thermal energy required to heat the aqueousprocess stream is supplied via heat exchange with a hot process stream,for example, via heat exchange with a catalyst recycle stream or aglycol absorbent stream.

In the distillation apparatus, the aqueous process stream is distilledand separated into an impurities stream, which is typically condensedand withdrawn as an overhead impurities stream, and a purified aqueousprocess stream, which is preferably withdrawn at or near the bottom ofthe distillation apparatus. The purified aqueous process streamcomprises water, optionally glycols, and a reduced amount of the one ormore impurities, as compared to the total amount of the impuritiespresent in the aqueous process stream supplied to the distillationapparatus. Thus, for example, if an aqueous process stream supplied tothe distillation apparatus comprises a total amount of 0.3 wt. % ofimpurities, the purified aqueous process stream withdrawn from thedistillation apparatus would comprise less than 0.3 wt. %.

At least a portion of the purified aqueous process stream withdrawn fromthe distillation apparatus is supplied to the absorber (e.g., forconstituting lean absorbent). If necessary, fresh water may also besupplied to the absorber.

A distillation apparatus suitable for use herein may comprise anydistillation apparatus known in the art for the separation and/orremoval of an impurity from an aqueous process stream. Morespecifically, a suitable distillation apparatus includes any device thatseparates water from at least a portion of the impurities present in anaqueous process stream based on their differences in volatilities byvaporization and subsequent condensation. Suitably, the distillationapparatus may separate water and at least a portion of the impuritiespresent in the aqueous process stream using one or more vapor-liquidequilibrium stages.

As will be appreciated by one skilled in the art, the design andoperation of the distillation apparatus can depend, at least in part, onthe type and concentration of impurities present in the aqueous processstream, as well as the desired composition (e.g., desired purity) of thepurified aqueous process stream. In some instances, for example, with abinary component feed, analytical methods such as the McCabe Thielemethod or the Fenske equation can be used to determine the number ofequilibrium stages to use to achieve the desired separation. For amulti-component feed stream, simulation models can be used for bothdesign (e.g., to determine the number of equilibrium stages needed inorder to achieve the desired separation) and operation (e.g., todetermine the optimum operating conditions). In addition, once thenumber of equilibrium stages is determined, one skilled in the art canuse known design techniques to readily determine the number ofseparation stages (e.g., the actual number of trays or height ofpacking) that can be used to achieve the desired separation. Typically,a distillation apparatus suitable for use in the present disclosure maybe operated in such a way as to include between 5 and 13 separationstages, more typically between 8 and 12 separation stages.

The distillation apparatus may comprise distillation trays (plates),packing, or a combination of distillation trays and packing. A packingis preferred because it may reduce residence time. Examples of suitabletypes of distillation trays include any type of plate commonly found indistillation columns, such as sieve plates, bubble-cap plates or valveplates, among others. The distance between each tray can besubstantially the same or alternatively, the distance between each traymay vary. In either configuration, the distance between each tray may beoptimized to allow for the best separation of the impurities from theaqueous process stream and/or to prevent entrainment between the trays.Additionally, in embodiments using packing, the packing material can berandom dumped packing such as, for example, Raschig rings, Pall rings,or Bialecki rings in metal or ceramic. The packing material can also bestructured sheet-metal packing.

In embodiments where packing is employed, the total required height ofpacking to provide the required number of separation stages can bedetermined by multiplying the number of calculated equilibrium stages bythe Height Equivalent to a Theoretical Plate, or HETP for that packing.The HETP is a value of the height of packing that will give the sameseparation as an equilibrium stage. As known to one skilled in the art,the HETP can vary depending on the type of packing selected. In someembodiments, the total height of packing can be split into one or morezones with vapor-liquid redistributors in between the zones, forexample, to accommodate height limitations due to packing structuralintegrity. In some embodiments, packing may offer the advantage of alower pressure drop as compared to trays, although consideration mustalso be given to the cost difference arising from the choice of traysversus packing.

The operating conditions within the distillation apparatus can beadjusted according to processing conditions. For example, thedistillation apparatus may be operated at a wide range of pressures,ranging from sub-atmospheric (i.e., vacuum), to near atmospheric, tosuper atmospheric. In practice, the general operating pressure of thedistillation apparatus can be selected during system design, althoughthere is some flexibility to adjust the pressure during normaloperation. The design operating pressure can range from about 60kilopascal (kPa) to about 220 kPa, preferably from about 80 kPa to about180 kPa, and more preferably from about 120 kPa to about 160 kPa.

The distillation apparatus may also be operated at a wide range oftemperatures. In practice, the operating temperature can be selectedduring system design, although there can be significant variation in thetemperature during operation. In some embodiments, there can be atemperature gradient present in the distillation apparatus, with thelowest temperature in the top portion and the highest temperature in thebottom portion. This gradient may be a gradual change across the columnand/or various sections of the column, or may be an abrupt temperaturechange. For example, at an operating pressure of 150 kPa, the operatingtemperature of the distillation apparatus may range from about 110° C.to about 113° C. As will be readily appreciated by one skilled in theart, the operating temperature and pressure of the distillationapparatus, and the composition of the aqueous process stream supplied tothe distillation apparatus, are interdependent.

The thermal energy needed for the operation of the distillationapparatus may be supplied by a heating mechanism placed internally orexternally to the distillation apparatus. For example, in a preferredembodiment, a reboiler may be employed. Optionally, the reboiler may beheated with steam or alternatively, the reboiler may be heated by heatintegration with a hot process stream, for example, a catalyst recyclestream or a glycol absorbent stream.

Preferably, the efficiency of removing the one or more impurities froman aqueous process stream is greater than 98%, more preferably greaterthan 99% and most preferably, greater than 99.5%. In addition,preferably, a purified aqueous process stream comprises less than 0.1wt. % impurities, more preferably less than 0.06 wt. % impurities, evenmore preferably less than 0.05 wt. % impurities, relative to the totalweight of the purified aqueous process stream.

Suitably, when an aqueous process stream comprises formaldehyde, theefficiency of removing formaldehyde from the aqueous process stream ispreferably greater than 30%, more preferably greater than 35% and mostpreferably greater than 39%. Similarly, when an aqueous process streamcomprises 2-chloroethanol, the efficiency of removing 2-chloroethanolfrom the aqueous process stream is preferably greater than 40%, morepreferably greater than 45% and most preferably greater than 50%.Further, when an aqueous process stream comprises one or more impuritiesselected from acetaldehyde, 2-chloromethyl-1,3-dioxolane,2-methyl-1,3-dioxolane and 1,4-dioxane, the efficiency of removing oneor more of these impurities from the aqueous process stream ispreferably greater than 98%, more preferably greater than 99%, and mostpreferably 100%.

In addition, preferably, a purified aqueous process stream comprisesless than 0.005 wt. % formaldehyde, more preferably less than 0.003 wt.% formaldehyde, even more preferably less than 0.002 wt. % formaldehyde,relative to the total weight of the purified aqueous process stream.Similarly, a purified aqueous process stream preferably comprises lessthan 0.002 wt. % of one or more impurities selected from acetaldehyde,2-chloromethyl-1,3-dioxolane 2-methyl-1,3-dioxolane, 2-chloroethanol and1,4-dioxane, more preferably less than 0.001 wt. %, even more preferably0 wt. %, relative to the total weight of the purified aqueous processstream.

By using the systems and processes disclosed herein, it is possible toreduce the amount of impurities present in an aqueous process stream,thereby allowing all or substantially all of the process water generatedin the production of ethylene glycol to be reused. The systems andprocesses disclosed herein provide the advantage that the volume ofwaste water generated in the production of ethylene glycol may bereduced, and further that impurities are not permitted to accumulate inthe process and have a deleterious effect on the quality of theresulting glycol product. Additionally, by carrying out these processesin an ethylene glycol manufacturing plant, it is possible tosignificantly reduce the amount of fresh water required, reduce theamount of waste water produced and reduce catalyst degradation, all ofwhich reduce operating costs.

Reference is now made to FIGS. 1 and 2, which are schematic views of aprocess and a reaction system for the production of ethylene carbonateand/or ethylene glycol, according to an embodiment of the presentdisclosure.

The reaction system shown in FIGS. 1 and 2 includes ethylene oxidereactor (2), which comprises an epoxidation catalyst. Epoxidation feedgas (1) is supplied to ethylene oxide reactor (2) via one or moreinlets, and typically comprises ethylene, oxygen, ballast gas (e.g.,methane or nitrogen), and a reaction modifier (e.g., monochloroethane,vinyl chloride or dichloroethane). In ethylene oxide reactor (2),ethylene is reacted with oxygen in the presence of an epoxidationcatalyst to yield ethylene oxide product stream (4). Ethylene oxideproduct stream (4) exits ethylene oxide reactor (2) via an outlet, suchas outlet (3), which is in fluid communication with a first inlet ofabsorber (6), such as inlet (5).

In absorber (6), the ethylene oxide product stream is brought intointimate contact with lean absorbent in the presence of one or morecarboxylation and hydrolysis catalysts. At least a portion of, andpreferably substantially all of, the ethylene oxide in the ethyleneoxide product stream is absorbed into the lean absorbent. Fat absorbentstream (8), which comprises ethylene carbonate and/or ethylene glycol,is withdrawn from absorber (6) via a first outlet, such as outlet (7),while any gases not absorbed in absorber (6) are withdrawn via a secondoutlet, such as outlet (9), and condensed to yield overhead absorberstream (10).

Overhead absorber stream (10) is supplied to vapor-liquid separator (12)(such as a knock-out vessel, flash vessel, etc.) via an inlet, such asinlet (11), to yield recycle gas stream (13) and aqueous bottoms stream(15). Typically, at least a portion of recycle gas stream (13) isrecycled back to ethylene oxide reactor (2), optionally after beingsupplied to a carbon dioxide absorption column and/or one or more guardbeds (not shown).

Aqueous bottoms stream (15) is withdrawn from vapor-liquid separator(12) via an outlet, such as outlet (14), and supplied as aqueous processstream (17) to distillation apparatus (19) via an inlet, such as inlet(18), which is typically located near the upper middle portion of thecolumn. Optionally, any portion of aqueous bottoms stream (15) thatbypasses distillation apparatus (19) may be combined with purifiedaqueous process stream (21) via line (16) and supplied to absorber (6)via a second inlet, such as inlet (23).

The thermal energy needed for the operation of distillation apparatus(19) may be supplied by any suitable heating mechanism, such as areboiler, and is preferably heated by using heat integration with a hotprocess stream, such as a catalyst recycle stream or a glycol absorbentstream. For example, as shown in FIG. 2, the thermal energy needed forthe operation of distillation apparatus (19) is supplied by reboiler(44), which is heated by heat integration with hot process stream (43)withdrawn from separator (39).

Before being supplied to distillation apparatus (19), aqueous processstream (17) is supplied to preheater (36), wherein the preheater (36) ispositioned at the above-mentioned minimum distance to the inlet (18) ofthe distillation apparatus (19) Optionally, as shown in FIG. 2, beforebeing supplied to preheater (36), aqueous process stream (17) is firstsupplied to flash vessel (37) to recover light ends (e.g., ethylene andmethane) as light ends stream (38), which is preferably recycled back toethylene oxide reactor (2) after compression. Preferably, preheater (36)is similarly heated by heat integration with hot process stream (43)withdrawn from separator (39), as shown in FIG. 2.

Overhead impurities stream (22) is withdrawn at or near the top portionof distillation apparatus (19) and is supplied to a sub-coolingcondenser (45) and is cooled to a temperature in the range of from 5 to95° C., and the resulting cooled overhead impurities stream (22) issplit into a reflux stream (22a) which is recycled to the distillationapparatus (19) and an overhead impurities stream (22b) which istypically disposed of as waste. Purified aqueous process stream (21) iswithdrawn from distillation apparatus (19) via an outlet, such as outlet(20), preferably located at or near the bottom portion of distillationapparatus (19), and is supplied to absorber (6) via inlet (23), forrecirculation as lean absorbent. Make-up water (24) can be supplied ifnecessary.

Fat absorbent stream (8), which comprises ethylene carbonate and/orethylene glycol, is optionally supplied to one or more finishingreactors, such as hydrolysis reactor (26), via an inlet, such as inlet(25) (e.g., to provide further conversion of any ethylene oxide and/orethylene carbonate that was not converted in the ethylene oxideabsorber). Finishing reactor product stream (28) is withdrawn from theone or more finishing reactors, such as hydrolysis reactor (26), via anoutlet, such as outlet (27) and supplied to dehydrator (30) as adehydrator feed stream via an inlet, such as inlet (29). In dehydrator(30), water is removed from the dehydrator feed stream to yielddehydrator bottom stream (31) comprising predominately MEG, and overheaddehydrator stream (33).

Overhead dehydrator stream (33) is withdrawn from dehydrator (30) via anoutlet, such as outlet (32). Optionally, all or a portion of overheaddehydrator stream (33) may be combined with purified aqueous processstream (21) via line (35) and supplied to absorber (6) via inlet (23).Similarly, all or a portion of overhead dehydrator stream (33) mayoptionally be combined with aqueous bottoms stream (15) and supplied asaqueous process stream (17) to distillation apparatus (19), via inlet(18). Further, all or a portion of overhead dehydrator stream (33) mayoptionally be combined with overhead impurities stream (22b) via line(34) and disposed of as waste water.

Dehydrator bottom stream (31) is withdrawn from dehydrator (30),typically at or near the bottom of dehydrator (30), and optionallysupplied to a glycol purification apparatus (not shown in FIG. 1) toseparate glycols and remove impurities. As shown in FIG. 2, dehydratorbottom stream (31) may optionally be supplied to separator (39). Glycolproduct stream (41) is withdrawn from separator (39), via an outlet,such as outlet (40), and is optionally supplied to a glycol purificationapparatus (not shown in FIG. 2) to remove impurities. Further, hotprocess stream (43), such as a catalyst recycle stream or a glycolabsorbent stream, is withdrawn from separator (39), via an outlet, suchas outlet (42), and is preferably cooled and recycled to absorber (6)via inlet (23). As previously mentioned, the thermal energy needed forreboiler (44) and preheater (36) is preferably supplied via heatintegration with hot process stream (43), such as a catalyst recyclestream or a glycol absorbent stream, withdrawn from separator (39).

We claim:
 1. Process for the production of ethylene carbonate and/orethylene glycol comprising: a) supplying an overhead absorber stream(10) withdrawn from an absorber (6) to a vapor-liquid separator (12) toyield an aqueous bottoms stream (15) and a recycle gas stream (13); b)supplying an aqueous process stream (17) comprising one or moreimpurities to a distillation apparatus (19) to yield an overheadimpurities stream (22) and a purified aqueous process stream (21),wherein the aqueous process stream (17) supplied to the distillationapparatus (19) comprises at least a portion of the aqueous bottomsstream (15) withdrawn from the vapor-liquid separator (12), wherein theoverhead impurities stream (22) is supplied to a condenser (45) and iscooled to a temperature in the range of from 5 to 95° C., wherein thecooled overhead impurities stream (22) is split into a reflux stream(22a) which is recycled to the distillation apparatus (19) and anoverhead impurities stream (22b); c) supplying at least a portion of thepurified aqueous process stream (21) and an ethylene oxide productstream (4) to the absorber; and d) contacting the ethylene oxide productstream (4) with the purified aqueous process stream (21) in the absorberin the presence of one or more carboxylation and hydrolysis catalysts toyield a fat absorbent stream (8) comprising ethylene carbonate and/orethylene glycol.
 2. Process according to claim 1, wherein the aqueousprocess stream (17) is supplied to a preheater (36) prior to supplyingthe aqueous process stream (17) to the distillation apparatus (19) andwherein the preheater (36) is positioned at a minimum distance to theinlet (18) of the distillation apparatus (19) and/or wherein thepreheater (36) has a minimum volume.
 3. Process according to claim 2,wherein the aqueous process stream (17) is supplied to a flash vessel(37) prior to supplying the aqueous process stream (17) to the preheater(36).
 4. Process according to claim 2, wherein the aqueous processstream (17) supplied to the preheater (36) additionally comprises atleast a portion of an overhead dehydrator stream (33) withdrawn from adehydrator (30).
 5. Process according to claim 1, wherein the one ormore impurities are selected from the group consisting of: formaldehyde,acetaldehyde, 2,3-epoxy-1,4-dioxane, 1,3-dioxolane, 1,4-dioxane,2-methyl-1,3-dioxolane, 2-methoxy ethanol, 2,2′-bis-1,3-dioxolane,2-chloromethyl-1,3-dioxolane, 2-chloroethanol, and a combinationthereof.
 6. Process according to claim 2, wherein the preheater (36) isheated via heat integration with a hot process stream (43) withdrawnfrom a separator (39).
 7. Process according to claim 1, wherein heat issupplied to the distillation apparatus (19) using a reboiler (44), andwherein the reboiler (44) is heated by heat integration with a hotprocess stream (43) withdrawn from a separator (39).
 8. Processaccording to claim 1, further comprising: e) supplying at least aportion of the fat absorbent stream (8) to one or more finishingreactors (26) to yield a finishing reactor product stream (28), and f)supplying at least a portion of the finishing reactor product stream(28) to a dehydrator (30) to yield a dehydrator bottom stream (31) andan overhead dehydrator stream (33).
 9. Process according to claim 8further comprising: g) supplying at least a portion of the dehydratorbottom stream (31) to a separator (39) to yield a glycol product stream(41) and a hot process stream (43); and h) supplying at least a portionof the hot process stream (43) to the absorber (6).
 10. System for theproduction of ethylene carbonate and/or ethylene glycol, wherein thesystem comprises: an absorber (6) comprising one or more carboxylationand hydrolysis catalysts, at least two inlets and at least two outlets,wherein a first inlet (5) of the absorber (6) is fluidly connected to anoutlet (3) of an ethylene oxide reactor (2); a vapor-liquid separator(12) comprising an inlet and an outlet, wherein the inlet (11) of thevapor-liquid separator (12) is fluidly connected to a first outlet (9)of the absorber (6); a distillation apparatus (19) comprising an inlet(18), an outlet (20) and an outlet (22), wherein the outlet (20) of thedistillation apparatus (19) is fluidly connected to a second inlet (23)of the absorber (6) and the outlet (22) of the distillation apparatus(19) is fluidly connected to a sub-cooling condenser (45); andoptionally a preheater (36) comprising an inlet and an outlet, whereinthe inlet of the preheater (36) is fluidly connected to the outlet (14)of the vapor-liquid separator (12) and wherein the outlet of thepreheater (36) is fluidly connected to the inlet (18) of thedistillation apparatus (19).
 11. System according to claim 10 furthercomprising: a flash vessel (37) comprising an inlet and an outlet,wherein the inlet of the flash vessel (37) is fluidly connected to theoutlet (14) of the vapor-liquid separator (12) and wherein the outlet ofthe flash vessel (37) is fluidly connected to the inlet of the preheater(36).